Two-stage hydrocarbon synthesis



Sept. 8,

Filed Dec. 18, 1948 C. E. HEMMINGER TWO-STAGE HYDROCARBON SYNTHESIS 2 .Sheets-Sheet 2 Patented Sept. 8, 1953 TWO-STAGE HYDROCARBON SYNTHESIS Charles E. Hemminger, Westfield, N. J., assignor to Standard Oil Development Company, a corporation of Delaware applicatif-m December 1s, 194s, serial No. 66,115

1 claim. (ci. 26o-449.6)

The present invention relates to improvements in the synthesis of hydrocarbons. More particularly, the present invention relates to a single or multi-stage hydrocarbon synthesis process employing cobalt as the catalyst.

-eretofore, a method of synthesizing hydrocarbons from a gaseous feed containing hydrogen and carbon monoxide kand employing a cobalt catalyst was known. This earlier process resulted in the production of a normally liquid hydrocarbon of high parainic content.

It has now been discovered that by employing a cobalt-type catalyst under conditions hereinafter set forth more fully, a hydrocarbon product including a gasoline fraction may be synthesized from carbon monoxide and hydrogen, which gasoline fraction will have an improved octane rating due to the fact that it contains substantial quantities of olenic hydrocarbons. In order further to improve the quality of the gasoline thus produced, the present invention contemplates aftertreating the gasoline with an active cracking catalyst such as a synthetic silica-alumina gel cracking catalyst or activated clays such as acid treated montmorillonite, or even activated bauxite.

An important feature involves the use of a relatively lcw hydrogen-carbon monoxide ratio in the feed gas, coupled with relatively high reaction temperatures.

It is a matter of record that heretofore relatively low hydrogen-carbon monoxidev ratios in the feed gas to the hydrocarbon synthesis process employing a cobalt-type catalyst had been proposed. However, the temperatures proposed in Vthese earlier processes were relatively low, for

the matter of controlling temperature Was difiicult since these proposals had to do with fixed bed type of operation.

It has now been discovered that by employing the fluid catalyst technique, low hydrogen to carbon monoxide ratios in the feed, that is to say, ratios within the range of from .'8 to 1.5 mols of hydrogen per mol of carbon monoxide (i. e., in the total feed) may be successfully employed with a cobalt-type catalyst, even at elevated temperature, without incurringthe diculties of the older fixed bed type of operation, where inordinately large quantities of methane and other normally gaseous hydrocarbonsl were formed.

It is another feature of this invention that although one starts with `cobalt catalyst of the Iproper size for good fluidization, after the operation has been on stream for-some period of time, there is a tendency-for the catalyst to undergo fragmentation to form fines. This invention includes the concept of permitting these nes to 'be removed from the reactorby elutriation. The

observable effect of this particular feature of the invention is that any tendency of whatever wax for heavy oils formed to interfere with good uidi- Azation of the catalyst is repressed.

Another feature of the present invention which aids in securing the foregoing desirable results is that there is a low conversion of carbon monoxide Aper pass, this conversion amounting to l0-50% of the carbon monoxide feed to the reaction. Of

course, in order to attain high ultimate conversions, the present invention contemplates either recycling the unconverted reactants and hydrocarbons lighter than butanes or passing them through a subsequent stage or stages.

An object of the present invention, therefore, is to synthesize hydrocarbons, using a cobalt-type catalyst to produce a normally liquid hydrocarbon product, includingVv a gasoline product of high octane rating, in good yields.

Another object of the present invention is to conduct the hydrocarbon synthesis operation employing the iiuid catalyst technique under conditions which will tend to maintain the catalyst throughout the extended period of operation in a state where it is adapted for good uidization.

Another object4 of the present invention'is to produce by synthesis using hydrogen and carbon monoxide as reactants in the cobalt-type catalyst, normally liquid hydrocarbons of increased oleiinicity.

Another object of the present invention is to synthesize hydrocarbons from hydrogen and carbon monoxide under conditions such that the product is substantially free of oxygenated hydrocarbons and, as such, is a desirable intermediate for the manufacture of valuable chemicals such as alcohols, fatty acids, lubricating oils, various detergents, etc., becauseof its high olefin Content. Y, K

Other and further objects ofthe present in- Y vention Willebe apparent from the following more detailed description and claims considered in f connection With the accompanying drawings.

cation of the invention may be'carried out, this particular embodimentemploying a single reactor andrecycle of unconverted reactants; and

In Figure 2 there is shown another modification of the invention in which two reactors-in series Y are employed with a Water gas shift between stages but no other adjustment of the feed.

Referring in detail to the process illustrated in Fig. 1, the operation about to be described is one in which the feed gas containing hydrogen and carbon monoxide is prepared from natural gas by partial oxidation with oxygen according to conventional methods. This process normally results in the production of a feed gas in which the hydrogen to carbon monoxide ratio is of the order of about 1.7 mols of hydrogen per mol of carbon monoxide. rThis then would be the molar ratio of hydrogen to carbon monoxide in the fresh feed. However, since this is 'a recycleoperation and since the recycle gaseous material has a hydrogen to carbon monoxide ratio of about 1 to 1.4 and When the recycle ratio is about 2, the ratio of hydrogen to carbon monoxide in the total feed is about 1.5. n

The fresh feed containing carbon monoxide and hydrogen enters the present system through line I and 'mixes with Yrecycled material from line 3 and thereafter, the mixture is discharged into line l from which it is forced vinto reactor 5 through a grid or other foraminous member G. Thereactor contains a body of iiuidized powdered cobalt-type catalyst C having an average particle size of from 30 to 60 microns. The catalyst is maintained in the form of a dense, turbulent fluidized mass by causing the gasiform reactants to flow therethrough at a superficial velocity of from 1/2 to 2 feet per second. By superficial velocity one means Vthe gas velocityat operating conditions of temperature and pressure were there'no catalyst in the'reactor. Sufficient catalyst is contained in the reactor so as to maintain an upper dense phase level at L. Above the Vlevel L the concentration of catalyst in the gasiform material decreases sharply and the exiting gasiform Ymaterial which is withdrawn through line 6 contains substantially only very fine catalyst such as catalyst having a particle ,size of less than 20 microns. At the beginning of the operation the amount of material entrained overhead with the exiting gases and vapors is relatively small but as time passes, the amount of such material carried overhead may be substantial. In order to compensate for catalyst lost in the exiting 'gasiform material, fresh catalyst may be'addedthrough valved line in order to keep the amount of catalyst in the reactor substantially constant.

The gasiform material in line 6 is forced through a solids-gas separating device 8 which may be, for instance, a centrifugal separator, and

in this separator the catalyst is substantially removed from the vapor and the thus recovered catalyst is withdrawn from the separator 8 through line S andaportion of this separated catalyst is returned to the reactor through Valved'line IB. The remainder of the catalyst separated in 8 is rejected from the system throughvalved line II.

VIt will be understood, of course, that While it is desirable to remove'fines from the reactor, nevertheless, -a certain quantity of the material having a particle size from v(l1 to 2O microns, should be maintained thereinin order to provide a fluidiza'ble mass. Hence, the material returned through line VIIJ contains fines and some entrained coarse material.

Without imposing a limitation thereon, Ybut Vmerely to illustrate the general order of particle "sizea'nd particle size distribution which will give a bed of powdered cobalt adapted to be fluidized readily, the following is set forth:

Percent by Particle Size in Microns Weight varound performance, such as good heat transfer to coolingsurfaces, constant bed Volume, high catalyst eiciency, or in other words, high yields of desired products per unit weight of catalyst, etc.

The material withdrawn through line II may be reworked to produce catalyst of the proper particle size for further use in the process.

The gasiform material which discharges into separator 8 is Withdrawn substantially freed of catalyst, through line I2, thence passed through cooler i3 where normally liquid materials are condensed, and thereafter, the product is discharged via line I4 into separation drum I5. The desired normally liquid product, after separation of water therefrom, is withdrawn fromdrum I5 through line It and delivered to a fractionator il where it is subjected to fractional distillation to recover the following products, namely 'an overhead fraction, which is Withdrawn Vthrough line IS, and contains naphtha and the lighter hydrocarbons. This material in line I8 is passed through a second cooler I9, and vthence discharged into a separation drum 2t, from which drum the normally gaseous material is Withdrawn overhead through line 2 I ,while the naphtha fraction is withdrawn through line 22, subjected (after removal of water) to treatment with a fluidized mass of powdered cracking catalyst such as silica alumina gels, etc., preferably one of low activity (D-I-L of 28, or less), for instance, a spent cracking catalyst. This treatment, which is for the purpose of upgrading the naphtha (improved anti-detonation quality), is carried out in a conventional catalytic cracking system consisting primarily of a cracking zone and a catalyst regeneration zone employing the fluid catalyst technique. Following the treatment of the naphthafraction in system 23, the product is Withdrawn through .line indicated at 25 and delivered to a finishing still 25, from which product gasoline is recovered through lineA 26, while the heavy bottoms are withdrawn through line 2. The latter form a minor quantity of the total product and may be used as a domestic heating fuel, since the treatment is rather mild since rich yields of treated product are, ofcourse, desired.

Referring again tofractionator column Il, the bottoms therefrom,Y comprising gas oil, withdrawn through line 28, may be subjected to catalytic cracking in a conventional cracking system or employed as a diesel fuel.

Referring againto separator I5, the gasiform material separated in separator VI5 is withdrawn overhead through line 29 and in-part, recycled through line 3 as previously indicated. lThe vmaterial not recycled'through line `3 file .y'lce delivered to a recovery vandpolymeri-zation' plant SI, in Whichthes and C4hydrocarbonsfarelseparated from theremainder of the `material and thence delivered to a polymerization plant where they are converted to polymer gasoline.

It is emphasized that with respect to the recycle stream, the same contains C1, C2, and Ca hydrocarbons and it has been noted that these are not merely diluents vin the reactor 1 when returned thereto, but are substantially converted to normally liquid oleiinic and aromatic hydrocarbons, which result is reiiected in the improved octane rating of vthe gasoline product. To further improve the product the same may be aftertreated as by contact with a cracking catalyst undercr'acking conditions as herein'disclosed.

Referring again to reactor l, in order to pur'- ify the catalyst and cleanse it of oily and waxy material, the same is withdrawn through a pipe 32 (aerated) continuously and discharged into branch line 33 containing some of the recycle material. Inline 33, theadhering oily and waxy material is substantially completely stripped from the catalyst and the mixture of catalyst and gasiform material is returned to the top of the reactor as indicated, the oily and waxy material exiting with the reaction products through line 6.

In Fig. 2, the operation is conducted as shown in the drawing in a two vessel system with a water gas shift step between the reactors.

To the accomplishment of this result, therefore, the feed enters the system shown through line and is then mixed with recycled gasiform material from. line |0| and discharged via feed inlet line |02 into reactor |04. Reactor |04 contains a fluidized mass of catalyst C1 which has an upper dense phase level at L1 and the gasiform feed is caused to flow through a distributing grid or the like G1. The eiiiuent vapors resulting from the reaction exit from the reactor through line |05. Since these vapors contain fines and coarse material entrained therein, they are forced through a solids-gas separating device |01 for the purpose of separating at least a portion of this catalyst material and returning it through line |08 in part, and in part Yrejecting it from the system through line |09. The gasiform vapors are withdrawn from separator |01, through line ||0, thence through a cooler where they are cooled sufficiently to liquify normally liquid products of the reaction including, of course, the water, the hydrocarbons vand any minor amount of oxygenated hydrocarbons. The liquied products are then conveyed by line I I2 into a separation drum ||3. From the bottom of separator ||3, through line A| |4, thenormally liquid products are Withdrawn and these may Vbe purified to recover desired products in a treater similar to 23 of Fig. 1.

The normally gaseous material is'withdrawn overhead from separator ||3, through line ||5, and is discharged into a water gas shifterA 6 where it is treated with steam with or without oxygen under known conditions -in order to convert the CO and steam to CO2 and` hydrogen, as well as to convert some methane to. CO rand hydrogen.

Alternatively, the gasiform material in line may be passed'via line ||1 into a conventional absorber H8 where it is treated with'a lean absorbent oil to remove .hydrocarbons .of yhigher molecular weight than ethane. The fat liquor from this operation may -be withdrawn lthrough line ||9, whereas the gas containing C1 :and .C2

hydrocarbons now substantially freed .of .heavier hydrocarbons, is'- withdrawnithrough 2 Bland' dis- 6 charged into water gas shifter I-B, through line 2|, cooled in cooler |22 and thence discharged via |23 into a CO2 scrubbing zone |24 containing a vconventional solvent for C02. The vapors substantially freed of lCQ?. are then withdrawn through line |25 and may be recycled in part through line |0| to line |02 and thence to reactor |04. The main portion Vof the vapors in line |25 is discharged through line |26 into the bottom of a second reactor |21. also containing a catalyst, C2, in the form of uidized bed. As usual, the iiuidized mass of catalyst forms an upper dense phase level at L2 and the reactor is provided with the conventional perforate distributing means G2.

The reactant vapors are withdrawn lfrom reactor |21 through line |28, andas before, are forced through gas-solid separating device |29 wherein fines `and coarser material are separated from the gasiform material and returned, in part, through line into reactor |21. Another portion of the withdrawn catalyst is rejected from the system through line |3|.

The vapors issuing from separator |29 are forced through line |32, thencethrough a-cooler |33 where normally liquid products are condensed, and thereafter the cooled material is forced through line |34 into a separator |345. The bottoms comprising .the liquid products are withdrawn from separator |35, through line |35 and these maybe combined with the liquid products of the rst stage in line v| |4 and delivered after separation of water to a iiuiditreater, vthat is, to a reactor or treater containing a ,fluidized mass of cracking catalyst, this'operation `being similar to that carried out-in23 -of Fig. -1.

The gasiform material is'withdrawn from Iseparator |35, through 'line |31 vand delivered to a conventional absorber |38 -wherein Vit is treated with a normally liquid lea-n 'hydrocarbon Voilfor the purpose of separating outhydrocarbons containing more than 2r-carbon atomsywhereupon the fat liquor containing these- 4absorbed hydrocarbone is withdrawn through line |39. The ,C1 and C2 hydrocarbons are withdrawn through line |40 Vand rejectedfrom vthe system or returned to the apparatus not shown `for production ofthe synthesis gas in linejl.

Referringv again to the lgasiform material in line |31, a portion ofthisfg'asiform material may 'be recycled to the reactor through line |4|.

As in the modification shown -in Fig. l, it is desirable for best operation that the catalyst kbe w-ithdrawn continuously lor intermittentlyatfrequent intervals to cleanse the same .by removing oilyv and waxy material therefrom, 'th-us from Yreactor |21 the catalyst ris withdrawn through line (aerated) and thence :discharged into |46 into thereactor1|21 :aboveitherdenseyphase level Lz and during its passagein*thisrtransfer line, the oily and waxy vmaterial .issubstantially stripped from the catalyst-and.exits,fromzthegre actor with the othersproductseof"thereaction :through line |28. l

Reactors 5, |04 and|21yare1providedfwith suitable cooling means zHi, 1H?. :and Hs respectively.V

These may takethe formsof coils lfimbeddediin the Vfluidized mass .through which :is circulated .a .suitable `iiowing coolant such E as: steam orf water.

An alternate method-,of operating ithegplant depicted in Fig.2is,=to,withdrawgfrcmrthe system, through line |50, excess'Ci and Ca hydrocarbons and unconverted CO and H2, and at the same time closing valve lill in; line 126 and directingall of the remaining gasiform material in line |25 to reactor |04.; In Vthis modification, it will be noted that although similar to the process described previously and illustrated in Fig. 1 in that a single vessel is employed with recycling, the present modification differs in that the gasiform material is shifted (CO-{II2O CO2+H2) prior to recycling to the reaction zone.

Having described several modifications of the invention while referring to the drawings, there are now set forth below operating conditions and temperature, pressure etc., which embody preferred modifications of the invention.

In the process illustrated in Fig. 1 the catalyst indicated is a cobalt type catalyst. It is important that the average size of the catalyst should be less than 60 microns, preferably as small as 30 microns average size, so that it is adapted for good fluidization in the reactor and the maximum surface of catalyst per unit per weight of catalytic material will be afforded to the reacting gases. In other words, the catalyst average particle size should be between 30 and 60 microns, which is a size range which pro- Vides a maximum surface per unit weight of catalyst, and at the same time, is adapted for good uidization.

Regarding the composition of this cobalt type catalyst, it should consist of cobalt metal precipitated on an inert carrier, such as kieselguhr, silica gel, alumina gels,'alundum, or other materials essentially made up of silica and/or alumina. There are many physical modifications of these materials which are obtained by different methods of drying and sintering. It is desirable that the modifications give the greatest mechanical strength to the carrying materials chosen. The preferred carrier is silica gel.

A good way to prepare the catalyst is to impregnate dry silica gel with, say, cobalt nitrate and soluble salts of promoters mentioned below, then treat withv ammonium hydroxide to precipitate cobalt hydroxide, thereafter dry and then reduce with a hydrogen-containing gas at a temperature of about 350 to 500 F. In addition, the catalyst should contain an activator, such as thoria, magnesia, manganese, and in addition, small portions of an alkaline metal compound, such as sodium carbonate etc. These may be included in the composition as indicated previously. A preferred composition of the catalyst is one with 100 parts by weight of cobalt, as metal, 4 parts of thoria, 4 parts of magnesia and 2100 parts of the inert carrier, e. g. silica gel. This catalyst may be promoted with sodium carbonate to give 1 to 2% NazO on the total catalyst composition. Upward of 10% by weight of iron may be incorporated in the catalyst with advantage, giving-a greater yield of olens.

In the fluid vessels described hereinbefore, it will be understood that the fluidizing reactants are forced through the reactor at a velocity suliicient to cause formation of the fluidized mass. Usually, this velocity (measured as though there were no catalyst in the reactor and at operating conditions, temperature and pressure) is of the order of t@ to 3 ft. per second.`

Conditions in reactor 5 Rango Preferred Temperature, Ff. 45o-600 525 Pressurap. s.'i 25-250 75 Feed Rate, C. F., COA-Hillb. C 2-50 l5 Recycle Ratio, Recycle/fresh feed. 1-5 2 lh/O O in total feedmf. 1. 2-1. 6 l. 5 Conversion (CCH-H1), Percent 5 90 Conditions in reactor 104 Range Preferred Temperature, 45o-eco 50o Pressure, p. s. i 25-250 75 Feed Rate, C. F., CO 2-50 10 Recycle Ratio, Recycle/fresh feed- 0.2-1.0 0.6 HQ COinto alfeed 0.8-1.4 1.1 Conversion (CO-FE1), Percent 50-70 60 Conditions in reactor 127 Range Preferred Temperature, F 45o-600 535 Pressure, p. s. i 25-250 75 Feed Rate, C. F., CO-k-Hq/lb. Cat 2-50 l0 Recycle Ratio, Recycle/fresh feed 0.1-0.8 0.3 Hz/CO in total feed 1.0-1.8 1. 7 Conversion (CD4-Hg), Percent 25-45 35 Treater 23 Range Preferred Temperature, DF 70D-950 900 Pressure, p. s. i 15-100 50 Feed Rate, lbs. oil/hr./1b. Cat 0.5-30 5 Conversion, Wt. percent 65+ to Ci- 3-10 7-8 As hereinbefore indicated the preferred ratio of hydrogen tocarbon monoxide in the total feed charged to the several reactors is the same range, 1.1 to 1.7 mols H2 per mol CO, and less than the normal total feed ratio of 2 mols H2 per mol CO for a cobalt catalyst. However, the fresh feed to the reactors is not the same and is governed by the source. Thus, in the modification shown in Fig. l, the fresh feed is prepared by reforming (methane-steam reaction), methane or natural gas, and the hydrogen to CO ratioof about 1.7 4is higher than where the fresh feed is prepared by gasifying coal or the like. In the modification shown in Fig. `2, where the fresh feedgas is made by gasifying coal with steam, since feed gas has a Hz/CO ratio of 0.8 to 1.4 and is deficient in hydrogen, it is necessary to subject it to a Water gas shift reaction to proportion the ratios of hydrogen to CO. In the case of the total feed to the reactors shown in Figs. 1 and 2, the ratio of hydrogen to CO is within the limits of .8 to 1.5, accomplished either by recycling unconverted material (Fig. 1) or shifting the unconverted material prior to recycling or charging to the second reactor (Fig. 2).

In order to further illustrate the invention, there is set forth below a specific example.

A run was made in a fiuid reactor at 75 p. s. i. pressure with a catalyst ground to 6% 0-20 mu, 5% 20-40 mu, 37% 40-80 mu and 52% 80+ mu size. It had` the following composition: 30.7% cobalt, 2.6% magnesia and 66.7% silica, the latter being a silica gel carrier. The fresh feed had V1.5-l.6 Hz/CO ratio and 3.1-3.5 parts of tail or recycle gas were recycled per part of fresh feed. The fresh feed rate was about 10 CF Hal-CO/ hr./lb. catalyst. The following results, which are were obtained:

These data show that increasing the temperature from 432 F. to 500 F. had the beneficial eiects of about doubling the conversion of fresh feed gas, increasing the bromine number, decreasing the aniline point and increasing substantially the percentage of gasoline in the collected oil sample without decreasing the selectivity or yield of liquid products per Volume of Hz-l-CO' converted. This is contrary to existing literature on hydrocarbon synthesis with cobalt type catalysts -which uniformily stated that synthesis at temperatures above 425 F. resulted in a large increase in the yields of methane and carbone dioxide. No change in the former is indicated by the similar yields of Ca-I- hydrocarbons and, since the Hz/CO consumption ratio was 2.0-2.2 in the 500 F. test, there was little or no production of carbon dioxide. These experiments were repeated several times with the foregoing results.

As previously indicated substantial quantities of fines are removed overhead from the reactors (e. g. reactor with the exiting reaction vapors by elutriation or entrainment. This is a desirable condition because these fines may become wetted with heavy oily or waxy material in the reactor if they are permitted to remain therein indefinitely and interfere with good fluidization by adhering to the walls of the reactor or to cooling surfaces, thus stagnating on such surfaces, eventually packing or solidifying to a hard mass.

The advantages of the present invention are:

(1) Maximum yields of gasoline and lighter hydrocarbons.

(2) High olen contents of the gasoline and lighter hydrocarbons.

(3) The attainment of motor gasoline from the processes described herein, having an A. S. T. M. of 70 to 75 or higher.

(4) The attainment of high CO conversion (95% or better) even though the ratio of hydrogen to CO in the total feed is relatively low.

(5) Minimizes problem of carbon deposits on catalyst, characteristic of the hydrocarbon synthesis using iron; which deposit weakens the catalyst and causes formation of inordinately large quantities of fines.

Numerous modications of the invention may be made by those familiar with this art without departing from the spirit thereof.

What is claimed is:

The method of synthesizing hydrocarbons, including normally liquid hydrocarbons which comprises charging a feed gas containing hydrogen and carbon monoxide in the ratio of about 0.9 to 1.5 mols of hydrogen per mol of carbon monoxide to a reaction zone containing a fluidized mass of cobalt type catalyst, maintaining a temperature within the reaction zone of about 450 to 500 F., permitting the reactants to remain resident in contact with the catalyst for a sufcient period of time to convert 40 to 50% of the carbon monoxide, withdrawing the reaction products from the reaction zone, condensing normally liquid constituents, separating uncondensed gasiform material, subjecting the said gasiform material to a water gas shift reaction, removing carbon dioxide from the thus treated gasiform material, charging the resulting gasiform material to a second reaction zone containing a fluidized mass of cobalt type catalyst, subjecting the reactants to synthesis conditions of temperature and pressure in said second reaction zone for a suicient period of time to effect the desired conversion, recovering a product from said second reaction zone, condensing normally liquid constituents contained in said product and recovering the desired products from both of said reaction zones.

CHARLES E. I-IEMMINGER.

References Cited in the le of this patent UNITED STATES PATENTS Number Namek Date 2,220,357 Steinschlager Nov. 5, 1940 2,224,048 Herbert Dec. 3, 1940 2,253,607 Boyd et al Aug. 26, 1941 2,271,259 Herbert Jan. 27, 1942 2,274,064 Howard Feb. 24, 1942 2,274,639 Scheinermann et al. Mar. 3, 1942 2,287,891 Linckh June 30, 1942 2,417,164 Huber Mar. 11, 1947 2,445,795 Millendorf July 27, 1948 2,447,505 Johnson Aug. 24, 1948 2,451,879 Scharmann Oct. 19, 1948 FOREIGN PATENTS Number Country Date 581,174 Great Britain Oct. 3, 1946 OTHER REFERENCES Anderson et al.: Fischer-Tropsch Synthesis Etc., vol. 39, No. 12, Ind. and Engineering Chemistry, pages 1548-1554. 

